Hydrocarbon conversion system



April 26, 1949. B. EvERlNG ET Al.

HYDROCARBON CONVERSION SYSTEM NNNNROMN 2 Sheets-Sheet 2 7B. l.. EVERING ETAL vHYDROQRBON CONVERSION SYSTEM mwN QN @MN April 26, 1949.

Filed Df.

Patented Apr. 26, 1949 HYDROCARBON CONVERSION SYSTEM Bernard L. Evering and Edmond L. dOuville, Chicago, Ill., assignors to Standard Oil Company, Chicago, Ill., a corporation of Indiana Application December 15, 1941, Serial No. 422,986

3 Claims. l

This invention relates to the isomerization of parai'inic hydrocarbons of the light naphtha boiling range by means of an aluminum chloridehydrocarbon complex in the presence of excess hydrogen chloride and excess hydrogen under pressure. This is a continuation-impart of our copending application Serial No. 308,480, led December 9, 1939, now Patent No. 2,443,606, which in turn is a continuation-impart of our earlier application Serial No. 245,570, filed December 14, 1938 (now U. S. Paten-t 2,266,012), and is also a continuation-in-part of our copending application Serial 361,022, led October 14, 1940 (new U. S. Patent 2,300,249).

Other investigators have proposed methods of producing isobutane and higher saturated branched-chain hydrocarbons from straightchain paraiiins using aluminum chloride as the catalyst, but these methods result in such loW yields of the desired products based on the catalyst consumed that they are much too expensive for practical use. Aluminum chloride in the presence of hydrogen chloride Very readily forms a complex with the hydrocarbons present, and the rapid degradation of this complex to an inactive sludge has been a major factor in the low yields and poor product distribution obtained by prior methods. We have found that excellent yields of the most desirable type of branched-chain saturated hydrocarbons can be obtained from normally liquid straight-chain and slightly branched-chain paran hydrocarbons by subjecting them to the action of an aluminum halide catalyst effective in causing the conversion of straight-chain to branched-chain parains in the presence of an excess of hydrogen chloride and under a relatively high hydrogen pressure. The hydrogen greatly retards the rate of deactivation of the catalyst, thereby allowing especially high yields of the desired products per unit weight of catalyst and reducing catalyst costs so that the process is economically attractive.

It is an object of our invention to provide a process for the production of branched-chain saturated hydrocarbons, particularly isopentanes and isomerized (particularly neohexane) hexanes, which process is characterized by high yields of the desired isomerization products and low catalyst consumption. Another object is to provide a process whereby light naphthas of W antiknock values are converted into products rich in saturated branched-chain paraffin hydrocarbons of high stability, high knock rating and of volatility suitable for use as airplane fuels or blending agents.

An important object of our invention is to provide a continuous commercial process for converting C5 and Cs saturated straight-chain or slightly branched-chain hydrocarbons into more valuable branched-chain hydrocarbons With minimum degradation of the charge to isobutane and gases and with maximum yields of maximum octane number products per unit Weight of catalyst employed. In other Words, our object is to so correlate the operating variables such as temperature, pressure, space velocity, catalyst concentration, activator concentration, hydrogen pressure, etc., for a low knock rating Cs-Cs parafiinic hydrocarbon charging stock so that at least 25 or 30 gallons of the Cs-C charging stock may be converted per pound of aluminum chloride employed to give upwards of 98% yields by volume of a gasoline fraction having a CFR-M knock rating at least about 10 to 15 units higher than that of the charging stock. No previous isomerization process of this type has been commercially successful because at least one or more of the operating Variables was not Within the optimum operating range or was not properly correlated with the remaining operating variables. The object of our invention is to define the optimum operating variables and to correlate said operating variables in such a Way as to provide a commercially successful isomerization process.

A further object is to produce a high knock rating blending stock for aviation fuel which can be blended in maximum amounts with commercial isooctane for obtaining an aviation fuel of desired Reid vapor pressure, volatility characteristics, heat content, stability, knock rating, etc. A further object is to provide an improved method and means for introducing materials into the system, recovering products from the system and utilizing by-product materials obtained in the system,

In the accompanying drawings which form a part of this specification:

Figure l is a schematic diagram of one type of apparatus for practicing the invention, and

Figure 2 is a flow diagram of a commercial plant for practicing the invention.

In copending application 308,480, our invention is described and claimed as applied to a wide variety of charging stocks including heptane and high end point naphthas. The present invention is specific to the correlation of optimum operating variables for a C5-Cs charging stock, i. e., one which is substantially free from heptane as well as aromatics and which is characterized by an end point within the approximate range of 145 to 180 F., preferably about 152 F. This charging stock will be free of heptane and aromatics by virtue of its boiling range and it should also be substantially free from oleiins. In other Words, our invention is not applicable to cracked naphthas or other charging stocks containing substantial amounts of olens unless all but a very small fraction of the olens are removed therefrom. The feed stock for our process pref- I erably contains about 50% of paraffin hydrocarbons and stocks containing at least 80% paraffin hydrocarbons are especially desirable. It should contain no heptane, certainly less than and preferably less than 5% thereof.

We have found that the optimum correlation of optimum operating conditions for this particular stock is substantially as follows:

Temperature: 200 to 350 F., preferably about rFotal operating pressure: 500 to 1500, preferably 850 to 900 pounds per square inch.

Fresh aluminum chloride requirement (none recovered) about 1/2% by weight based on hydrocarbon charge, i. e., about one or two pounds While the materials introduced into our isomerization system are simply charging stock, aluminum chloride, hydrogen chloride and hy drogen, the actual catalyst in the system is an aluminum chloride-hydrocarbon complex. Butane does not form such a complex and while butane may be included in our charging stock our process is primarily for the isomerization of C5 and Cs hydrocarbons and it should not be confused with the very different aluminum chloride process for isomerizing normal butane. The aluminum chloride-hydrocarbon complex When freshly prepared may have a specific gravity as low as about 1.2. On continued use its specific gravity increases to a value within the range of about 1.6 to 2.0, its average specific gravity may be about 1.5. This complex is a fluid which is relatively non-viscous in its fresh state but becomes more and more viscous with continued use. The specific gravity of C5-C6 hydrocarbons is only about 0.66. Therefore, when introduced at the base of a continuous reactor along with hydrogen chloride, the hydrocarbons pass upwardly through the complex as a dispersed phase. If hydrogen is introduced at the top of the reactor a stirrer may be required, but when hydrogen is introduced at the base of the reactor no mechanical stirring means are necessary.

It is difficult in a continuous process of this type to determine the actual time of contact be tween oil and catalyst. In batch processes the severity of treatment at any temperature is indicated by a contact factor which may be expressed as:

Where C=weight of catalyst in reaction zone O=Weight of liquid hydrocarbons in reaction zone, and

t=average time of contact between hydrocarbon feed and catalyst.

In our continuous process it is somewhat more convenient to define the contact factor in terms of liquid feed rate according to the following equation:

where C`=weight of catalyst in reaction zone, and

F=rate at which liquid hydrocarbons are charged to the reaction zone in weight units per unit time The factor K has the dimension of time and will hereinafter be expressed in minutes.

In continuous catalytic processes it is often more convenient to express the severity of treatment at any temperature in terms of space velocity, i. e., the volumes of liquid feed charged per hour to the reactor per volume of catalyst which is retained in the reactor. The space velocity in our system may be ascertained from the above formulae since the specific gravity of the average catalyst complex is about 1.5 and the specific gravity of the charging stock is about .66. Equation 2 thus becomes:

q 1.5Vc 9V. (3) KF .66V/min.4V0/min.

where Vc is the volume of catalyst complex in the reactor, and

Vo is volume of oil charged to the reactor in corresponding volume units.

Transposing Equation 3 We have:

Since Vo/min. is 1/60 Vo/hr. we have:

Transposing Equation 5 we have:

(6) V"/hr'=1g=space velocity based on complex the contact factor K is minutes the corresponding space velocity is 27 Vo/hr./Vi-l based on complex or 44 based on the aluminum chloride content of the complex. Where the contact factor is 63 minutes the corresponding space velocity is about 2.2 Vo/hr./Vc based on complex or 3.5 based on aluminum chloride in the complex.

From Equation 6 it will be seen that space velocity varies with the reciprocal of the contact factor, i. e., a minimum contact factor cor responds to a maximum space velocity. The minimum value of the contact factor K that can be employed in carrying out the invention advantageously varies with temperature approximately as indicated by the equation:

The maximum contact factor in the case of very LogmKm 6.417

old catalyst may reach values of the order of one thousand times the minimum but for commercial operations the optimum contact factor is about ten to two hundred, usually about twenty to one hundred times the minimum. The above optimum conditions are for tower reactors;

for batch reactors the optimum contact factor 'l catalyst tends to decrease with time but is maintained substantially lconstant by the addition of fresh aluminum chloride and the withdrawal of spent complex from the system. The holding time of the catalyst in the reactor may range from about a week to several months. Only about .1 to 4 pounds, i. e. in the general vicinity of one pound, of a-luminum chloride is introduced per barrel of stock charged and for each barrel of stock ycharged per hour there is about 300 pounds of aluminum chloride in the reactor.

While temperatures of the order of 200 to 350 F. may be employed in our process we have found that the best results Iare obtainable at temperatures `of the order of 250 F. The use of temperatures below 300 F. results in at least three important advantages: First, and most important, the equilibrium oct-ane number in this particular temperature range is considerably higher than at temperatures of the order of 300 to 400 F. which means that we can produce a `product having an octane number lof about 84 at about 250 F. instead of only -about 81 at :temperatures of about 330 F. using the most favorable operating conditions for each case. In other words, the reaction is displaced in the direction of the more valuable isomers. The next advantage of the lower temperature range is that cracking is minimized so that there will be less :production of gaseous hydrocarbons and a greater yield of liquid Vproducts. The third advantage of the low temperature treating range from 200 to 300 F. lies in the fact that catalyst activity is actually longer in this range than at higher temperatures so that less hydrogen need be consumed or so that more conversion may be obtained with an equal amount of hydrogen.

The amount of aluminum `chloride employed in our process is vitally dependent upon other operating conditions and particularly upon the hydrogen pressure under which the reaction is ef- Table I Maximum Space Optimum Space l Velocity Velocity Temperature mhlk mglk AlCl in AlClgn Complex Complex Complex Complex Min. Min.

0. 5 5-100 273 444 1. 4-27 2. 2-44 l. 0 10-200 137 222 0. 7-14 1. l-22 3. 0 30-600 46 74 0. 23-4. 6 0. 4-7 10 10D-2000 14 22 0. 07-1. 4 0. 1-2

From the above table it will be seen that a space fected. Table II illustrates the `process carried velocity of about 1.4 may be used at temperatures from 200 to 330 F. although slightly higher space velocities should be employed in the higher temperature range and slightly lower space velocities in the lower temperature range. In the appended claims the expression a space velocity of about 1.4 volumes of charging stock per hour per volume of catalyst material in said reaction zone is hereby dened to be the optimum space velocities substantially as indicated by the Aabove Table I, i. e., said expression is hereby dened to include about 1.4 to 27 at 330 F., .2 to 4.6 at 250 F., .07 to 1.4 at 200 F. and corresponding ranges at other temperatures in the optimum temperature range.

Roughly speaking, this means :that for each barrel per hour of stock charged there should be in the general vicinity of about 400 pounds of complex containing about 300 pounds of aluminum chloride. in the reactor. The activity of the out under a hydrogen pressure of 600 pounds per square inch for a l"total run length of 14.78 hours and in the presence of about 3% of hydrogen chloride by weight based on charge. In this run one part by weight of aluminum chloride effected the conversion of parts by weight of charging stock to give a 97.7% yield of an 81 octane number blending stock. The charging stock in this case was a C5C6 parafinic naphtha having a Iboiling range vof 1l0-l53 F. and having a CFR-M octane number of 67.5. The last column of :this table illustrates the continued activity of the catalyst even after long periods on stream, and shows that the catalyst was not spent at the end of vthe run. When catalyst activity is maintained by adding 1 or 2 pounds of fresh aluminum chloride per` barrel of stock charged and withdrawing equivalent amounts of fresh catalyst, .the catalyst life or holding time in the reactor is about l or 2 weeks.

about 200 cubic feet per barrel of stock charged at a temperature of about 250 F., a pressure of about 850 pounds per square inch and a hydrogen chloride concentration of about 3 to 8% by weight based on stock charged. About 100 to 300, preferably about 290 cubic feet of hydrogen per barrel of stock charged provides the necessary amount for the reaction and also provides the necessary agitation in the reactor so that mechanical stirrers are not necessary.

The hydrogen chloride charged to the reactor may vary from about 3% to about 8% or more based on naphtha charged. The use of the higher concentrations permits the use of somewhat higher space velocities. For example increasing the hydrogen chloride concentration from 2.8 to 4.4% Will permit an increase in space velocity of about 40%.

Our invention will new be described in more detail in connection with the apparatus shown in Figure 1. The normally liquid feed, such as alight, 158 F. end point naphtha containing normal pentanes and hexanes, is introduced into the system by means of pump Iii and line H and thence into the lower portion of the reaction vessel I2 which is shown as a jacketed pressure vessel equipped with a stirring device I3 so that the reaction materials are thoroughly contacted. The desired reaction temperature is maintained by passing a suitable gaseous or liquid heating agent through the jacket I4 of reaction vessel I2 by means of inlet I5 and outlet IE. In high temperature operation, saturated hydrocarbon gas consisting predominantly of propane and/or at least one of the butanes, catalyst slurry and activator are introduced into line il and mixed with the feed therein by means of pumps I'I, I8 and I9, and lines 20, 2I and 22 respectively. When the preferred low temperature operation is employed, pump I! is not necessarily used.

Free hydrogen is supplied to the upper portion of reaction vessel I'2 through pump 23 and line 24, and is there maintained at the desired reaction pressure, which is suiciently high to cause the hydrogen to dissolve in the agitated reaction mixture at a rate at least as great as it is used up in the reaction. Obviously if desired a number of reaction vessels can be used in series or parallel in place of the one shown, or vessels of other types well-known in the art can be substituted therefor.

A portion of the entire reaction mixture is continuously withdrawn from the upper portion of vessel i 2 through line 25 and passes either through valve 26 and cooler 2'! or through bypass valve 28 or partly through each valve into separator 29. The products consist of a catalyst complex which settles out in the lower portion of separator 25, and an upper layer consisting of a mixture of hydrocarbons containing branched-chain parafns having from 4 to 6 carbon atoms per molecule` unreacted feed stock, dissolved hydrogen, HC1 and unreacted parafnic gases if such have been charged. The catalyst complex is continuously withdrawn 'from separator 29 through line 3i! and either recycled to line 2l through valve 3i, line 32 and pump 33 or withdrawn from the system through valve 34 and line 35, or under Vsome conditions a portion of the complex may be continuously withdrawn from the system and the remainder recycled.'

The substantially spent complex can, of course, be regenerated or the aliuninum halide recovered therefrom and reintroduced into the system through pump I8. Furthermore, at least a portion of the spent complex can be treated with water or otherwise to furnish hydrogen halide for use as activator in the process.

The upper layer is removed from separator 29 through line and valve 31 and passed through valve 3B into fractionating tower 39, valves 4I), 4I and 42 in lines $3, 44 and 45, respectively being closed. Valve 38 is preferably of the pressurereducing type adjusted to the desired fractionating pressure. Fractionating tower 3Q is of a conventional type provided with two sidestream outlets 45 and el and is operated so that the bottoms therefrom contain undesirably heavy hydrocarbons, the normally liquid hydrocarbons falling within a desired boiling range are withdrawn through outlets 46 and lll and gases vhaving less than 5 carbon atoms per molecule are withdrawn overhead through line 43. The heavy liquids collecting at the bottom of fractionator 39 are withdrawn through line #i9 and may be recycled to line I l for further treatment through valve 5?, line 5I, pump 52, and line 53. Under some conditions it may be desirable to withdraw these heavy liquids from the system and this can be done through valve and line 55.

The sidestreams consisting predominatly of branched-chain paraffin hydrocarbons withdrawn through lines 46 and 41 are sent to storage by means of valves 55 and 51 and lines 58 and 59, respectively, valves 5i! and SI being closed. By thus keeping the desired products separated into relatively light and relatively heavy fractions, their use as blending constituents for m0- tor fuels is facilitated and stabilization if necessary can be carried out only on the light product. However, by closing valve 51 an-d opening valve 6! the entire product can be withdrawn in a single stream through line 58.

The overhead passing through line 48 consists of excess hydrogen propane, isobutane and possibly some normal butane, and also hydrogen halide and this overhead is preferably recycled to line 2l! through valve B2, line 63, pump et'. and valve to inhibit the conversion of the feed into such gases and reduce the quantity of the various gases which must be introduced from outside the system. During this procedure, of course, valves 65 and 61 leading to a further fractionation system, are closed and valve 6.8 controlling a vent is at least partly closed. In the event impure hydrogen is used the system must be purged of inert gases, either intermittently or continuously, for example, through valve 53 or a valved vent 69 on line 45. If it is desired, however. to recover the isobutane formed during the process, valve 62 is closed and valves 56 and 5'! are opened so that the gas stream passes through cooler 1t, pump 'H and line 'I2 into fractionating tower 73 which is operated under such conditions that the liquid bottoms contain the hydrocarbons having i carbon atoms per molecule and the overhead which is withdrawn through line 'M and valve 61 for recycling as described consists essentially of propane and hydrogen and also' hydrogen halide. The C4 fraction is withdrawn from the bottom of tower I3 through line 'i5 and consists predominantly of isobutane formed from the liquid feed.

Obviously the overhead passing through line e8 during low temperature operation will contain substantially no propane, and may contain very little isobutane. However, isobutance will be present in considerable quantities if the reaction conditions other than temperature are relatively severe, or if normal butane or isobutane 11 is'7 charged to reaction vessel I2r and. it can be separated from. the hydrogen and other light gases in fractionating tower l3- as described above.

A. variant of the' above-described procedure which is often advantageous isy carried out by reducing the pressure in separator 29 by manipulation of valves and 28, thus causing a gaseous phase. to form therein consisting pri marily of hydrogen, hydrogen halide and normally gaseous parans. This mixture of gases can be. recycled through valve l2 and line 452 to line 63 and pump 64,. and/or to the. inlet of pump 23, Whilethe upper liquid layer is passed. to fractionating tower 3S by means of valve (il and line 44. One advantage of this arrangement. is that the volume of gases in tower 39 and compression costsY are reduced, and better fractionation isobtained.

In another type of operation which is advan L tageous if it is desired to obtaina product having on the average a larger number or side chains and therefore a higher antiknock value, valve 4.0 is opened during the. early stages of a runso that most of the reaction products are recycled through lines 4:3 and 5l, pump 52 and line 53.. The branched-chain paran-'msupon passing again through the system tend to be come.4 more branched in. conguration and. con.- sequently have a still higher antilrrioclry Value. More and more of the products flowing through line 36 are thenallowed to pass through valve 38 tol the. fractionating tower 39 in which these branched-chain hydrocarbons are recovered as described.-` aboye. a` certain. percentage or the totalproducts, however, continuing toreturn throughvalve 40 to: the reaction, chamber 1.2;.

Another method of accomplishing substantially the same: result consists in. opening valves 60 and 50, rather than val-Ve 40. and recycling the relatively heavy sidestrearnproduct withdrawn from fractionating tower 39 through line, Il@ and it may even be desirable in the early stages to recycle the products Withdrawn fromv tower 39 through line 41 by opening valve 5I. method, however, flow through valve Ell is gradually restricted sothatonly a part of the products isi-recycled.

Still another method. of operation which is. ap'- plicable when isobu-.tane is desired asA a principal product is to close valve 5,6 entirely recycle the. entire heavy liquid product to be bro-kendown into isobutane, which action. can be` facilitated by using relatively large amounts ofV catalyst and introducing little or no butano or isobu-tane into f.

the system through pump: Il and line 20. Bry closing Valye 5l and opening valve 6i. the lig-ht liquid product, can be similarly recycled. Inthis method of opera-tion val-ye 62 is.. of course, kept closed and valves 66' andl-l.- leading tothe isobu.- tane recovery system areopen..

In Figure 2l We have. shown a commercial design for a 109.9. barrel per day isomerization. plant. About 5000. barrels per day of light naphtha are charged by pum-p1=6.through.heat exchanger.K H- to intermediatepoi-nt of light naphtha fractionator 'i8 which may be about 41/2 feet in diameter by to 55 feet high. This fractionator is provided with conventional heatingmeans 't9' at its. base and the fractionator is operated under such conditions that pentanes and heXan-cs are taken overhead While' heptanes and heavier hydroca-rbons are withdrawn from the base of the column through line 8B. rllhe hexanes., pcntarles and any butanes that. may be presentare taken As in the previous overhead through. line.- 8I, through cooler 82 to reflux drum 83. A portion of the. reilux condensate is returned by pump 84 through line 85 to the top of the fractionator to serve as reflux. When a substantially butano-free charging, stock is desired, any number of fractionating colunms or strippersl may be employed to obtain a. charging stock consisting of. pentane.. hexanes, or any desired. mixture thereof. In the. drawing this is diagrammatically illustrated. by the Withdrawal of a side stream pump 81 through. cooler 88.. In this case the rest. of. the reflux condensate. from drum 813 may be: withdrawn from the4 system through line. 8.5. About 9.0% or more of.v the side stream from. column. 18. may be introduced through line 89 to the. top ofA absorber llll and the rest of the stream may be introduced through line 9i to aluminum chloride slurry tank. 3L UsuallyI We prefer to leave'. any butanes. in. the charging stock so that, the side stream draw-off is. unnecessary.. Thus in our preferred. example, about. 900 barrels. per day or. more of reflux condensate is. introduced through line 3.3. to the top of. the absorber and about. 10D barrels per day or less ot the reflux condensate. is introduced through line. 9.4 to slurry tank 92..

The hydrogenchloride required for the reaction is absorbed in the. major portion of' the. feed stock before it is admixed with the aluminum chloride andv introduced into the. reactor. The hydrogen. chloride absorber may be about, l le, feet in diameter by 28 feet high. A stream oi hydrogen chloride gases from the system is introduced at the base, of this absorber through line a5. Make-up hydrogen chloride may be introduced through.. line 96y from apressure cylinder or by means of. a compressor. Instead of make-upY hydrogen chloride we. may employ chlorine, an alkyl chloride or other substance which will supply the necessary halogen halide activator under reaction conditions. We prefer, however, to employ hydrogen1 chloride and, to generate itA if necessary in a separate generator.

A. hydrogen chloride generator 9J may be of any known type- The chlorine. supplying agent introduced through line 98 is preferablyI chlorine gas. although it. may be sodium chloride or other halogen contain-ing reagent. The hydrogen supplying agentk introduced through; line 99 may be hydrogen gas.. aY hydrocarbon, sulfuric. acid, etc. Thus hydrogen and chlorine may be burned in generator 91 to supply hydrogen chloride; Wax tailings or other hydrocarbons; may be introduced through line 99 and chlorinated by chlorine gas introduced by line 98` to produce hydrogen chloride and chlorinated hydrocarbons (additional hydrogen chloride may, of course, be obtained from the latter). Sodium chloride may beV introduced. through line 98 and sulfuric acid through line 99. No invention is claimed in the specific means for generating hydrogen chloride. but in ourv system this hydrogen chloride generator op.- erates under such pressure. that. no compressors are required for introducing thev hydrogen chloride through line; l tothe. base or absorber and hydrogen chloride does not require the purification which is. generallyA necessary even for the production of commercial grades of hydro'i chloric acid. Biz-productsI from the hydrogen chloride generator are withdrawn. through line HH Instead of introducing the make-up hydrogen chloride under pressure directly into.vv absorber Sill, We may generate it at about atmospheric pressure, absorb it in all. or in a portion. of the feed stock. at low' pressurei and then pump the solu- `tion up to the necessary pressure for introduction into absorber 90 or into the reaction chamber.

The hydrogen chloride picked up in absorber 90 should be sunicient to give an amount of hydrogen chloride in the stock entering the reactors within the approximate range of 2% to 10%. preferably about by weight based on stock charged. From about 1,/2 to about 273 or more of this hydrogen chloride is obtained by gases introduced through line 95. Unabsorbed gases such as small amounts of hydrogen, methane, ethane, etc. are purged from the system through line |02, thus eliminating not only gaseous impurities from line 95 but also gaseous impurities from line |00. The hydrogen chloride rich charging stock from the base of absorber 90 is pumped by pump |03 through heater |04, through lines |05 and |06 to the base of rst reactor |01 at a pressure within the approximate range of 500 to 1500 pounds per square inch, perferably 850 or 1000 .Y

pounds per square inch. Hydrogen from source |08 is introduced by compressor |09 and line ||0 into line |05 in amounts within the approximate range of 100 to 300, preferably about 200 cubic feet per barrel of stock charged to the reactor .l

(the hydrogen being measured under standard conditions) Aluminum chloride from source is introduced through line or hopper H2 into slurry tank 92 at such a rate that the amount of aluminum chloride introduced through line |05 by pump ||3 and line 4 is about 1 or 2 pounds per barrel of charging stock introduced therethrough. The stream entering reactor |01 through line |06 is at a temperature within the approximate range of 200 F. to 350 F., preferably about 250 F.

Reactor |01 may be a vertical tower about 5 or 6 feet in diameter lby about 18 or 20 feet tall. When the reaction is initiated the major part of this reactor may be lilled with an aluminum chloride-hydrocarbon complex the density of which is within the approximate range of 1.3 to 1.7 but .which may be maintained at about 1.5 by methods Vhereinafter described. The charging stock, therefore, passes upwardly through the complex while intimately dispersed therein. Suitable baille plates or packing may, of course, be employed to prevent channelling. The space velocity under the recited conditions is approximately onehalf volume of liquid charging stock per hour passes by line I8 through cooler ||9 and pressure reducing valve so that it enters the cool settling chamber |2| at about 100 F. or less and at a pressure of about 250 pounds per square inch. The cool settler may be a horizontal or slightly inclined drum about 5 or 6 feet in diameter and about 16 feet in length. Released gases leave the top of settler |2| through line |22 which discharges into line 35. Aluminum chloride and catalyst material is thrown out of solution in the cool settler and the clear product liquid is withdrawn from an upper point in this settler through line |23 and introduced into hydrogen chloride stripper |24.

Stripper |24 may be a column about 3 feet in 14 diameter by about 33 feet high and may be provided with heating means |25 at its base. It may be operated with a top temperature within the range of about 100 to 150 F. and a bottom temperature within the range of about 300 to 400 F. The removed hydrogen chloride together with released gases such as hydrogen, methane, ethane, etc. is taken overhead through line |26 to line 95. The liquid from the base of the stripper is introduced at the base of scrubbing tower |21 either directly through line |28 or through a cooler |29.

The scrubber |21.may be a tower about 4 feet in diameter by about 32 feet high and may be provided with suitable bales, trays or bubble plates for effecting intimate contact of the upflowing products with caustic introduced through line and water introduced through line 3|. The product is neutralized in the base of this scrubber and is water washed in the upper part thereof. Spent caustic solution is withdrawn from the base of the scrubber through line |32.

The water Washed product passes from the top of scrubber |21 through line |33 and heat exchanger |34 to an intermediate point in stabilizer |35. This stabilizer is provided with conventional heating means |35 at its base. Butanes and any lighter products are taken overhead through line |31, through cooler |38 to reux drum |39 from which gases may be vented through line |40. Condensed reflux may be returned by pump |4| and line |42 to the top of stabilizer |35. The hydrocarbon stream consisting chiey of isobutane may be withdrawn from the system through line |43.

If desired, a single isomate fraction may be withdrawn from the base of the stabilizer through line |44. We may, however, withdraw only the heaviest isomate at this point and we may withdraw a light isomate as a side stream through line |45. Here again it should be understood that in actual practice a plurality of columns or towers will be employed for effecting any desired fractionation of isomate. The isomate consists essentially of isohexanes and isopentanes and it may be fractionated to insure the removal of any unconverted hexane or higher boiling products which may be formed and to obtain a product of desired Reid vapor pressure for blending in desired amounts with isooctane to make a super aviation fuel.

Returning to the reaction system we may employ a second reactor |46 of about the same size as the rst reactor |01. Products from the rst reactor, instead of going to the settler through line H6, may pass through line |41, heat eX- changer |48 and line |49 into the base of this second reactor. The operating conditions in the second reactor are substantially the same as in the rst reactor although it may be operated at somewhat lower temperature. Products from the top of the second reactor pass through line and line ||6 to Warm settler ||1 as hereinabove described.

Instead of operating the reactors in series they may be operated in parallel by passing only a part of the charging stock through line |00 in the rst reactor and by passing the remainder of the charging stock through line |5| and line |49 to the base of the second reactor. By means of this arrangement one reactor may be on-stream while another reactor is standing by for repair or replacement of catalyst complex.

During the course of the reaction the catalyst complex tends to become less active. A substantag-468,549

tia-lly constant catalyst activity may be maintained in the reactors by continuously withdrawing a portion Yof the catalyst from the base of the reactors at about the same rate as additional complex is formed with the aluminum chloride and Ihydrogen .chloride 'intoduced with the charging stock. Thus catalyst from the second reactor may be withdrawn through line |52 by means of pump |53 and either introduced. through line |55 to the rst reactor, Withdrawn from the system through :line |55, or introduced through line |55 to hydrogen chloride recovery drum |51. Catalyst from the base of the first reactor may be withdrawn through line |58 by means of pump |59 and introduced through lines |60 and ||5| to the second reactor, or introduced to the hydrogen chloride recovery drum |51 through line |62, or withdrawn from the system through line |83. If the second reactor `operates .at a lower tempera- `ture than the iirst reactor and in series therewith we prefer to introduce catalyst from the second reactor through line |54 to the iirstreactor and to remove catalyst through lines I6@ and |62 from the first reactor to the hydrogen chloride recovery drum. For such operation we prefer to introy lduce a part or all of the make-up aluminum chloride slurry from line ||4 to 'the second reactor A to the second reactor yor lin-e |56 to the rst reactor.

The spent sludge may be either discarded .through the system through lines |55 and |53 or utilized in any conventional manner. We prefer, however, to introduce the spent sludge into drum |51 and to add to the sludge i-n this drum through line |69 a sumcient amount of Water or sulfuric acid to eect recovery oi the maximum possible amount oi anhydrous hydrogen chloride. The recovered hydrogen chloride is returned by line Vit to line 95 and the coke or residue is Withdrawn vfrom drum i? through line ill, If water is employed it should be used in less than stoichiometrc amounts in order thatthe recoveredhydro- :gen chloride be substantially anhydrous; sludge will thereupon be converted into a cokey mass which may be lremoved from the drum by hydraulic or rany other conventional decoking means. A larger amount of hydrogen chloride may .be recovered by the introduction of sulf densed hydrocarbons, chiefly butanes, may be Awithdrawn from the base of reilux drum |14 by `means of pump |16. A part of these liquefied 'hydrocarbons are introduced through line |11 `to the top orf tower |26 in order to'provide reilux liquid. Another part of -'the reflux condensate may be withdrawn from the system through line |18. If the Withdrawn fraction is to be' employed the in an. aluminium chloride alkylation process it will he unnecessary to remove any dissolved hydrogen chloride therefrom otherwise this stream i ay, of course, be neutralized with caustic and Vd in a suitable vwashed tower.

of the features of our `invention is the solution of the hydrogen chloride in at least a major part Vof the charging stock before this stock is introduced into the reactor. We have found that the solubility of hydrogen chloride in this light naphtha charging stock markedly increases with increase in pressureffhus at F. and 30 pounds per square inch hydrogen chloride pressure only about 1.5% oi Vhydrogen chloride will dissolve in the liquid feed but at 100 F. and 100 pounds I'per square inch hydrogen chloride pressure about 5% of hydrogen chloride can be dissolved in said naphtha. At 60 F. and 1Go pounds per square inch hydrogen chloride pressure 9 or 10% by weight of hydrogen chloride will dissolve vin the light naphtha. At 60 F, and 290 pounds per square inch over 20% of hydrogen chloride Will dissolve in the light naphtha. The temperature and pressure in the hydrogen chloride absorber should be such as to effect substantially complete solution of the hydrogen chloride out of the gases which pass through this absorber countercurrent to the upfiowing light naphtha stream. We prefer to operate this absorber at a pressure Within the approximate range of 50 to 400, preferably about 10o to 300,A pounds per square inch and at a temperature readily obtainable by ordinary cooling water, namely, from room temper- 'ature to upwards of 100 F.

Once the hydrogen chloride is in solution and this solution is subjected to reaction pressure by pump |63, the bulk of the hydrogen 4chloride is kept in solution even at reaction temperatures. In iact, 5% of hydrogen chloride will dissolve in the light naptha at a, temperature of 250 and a pressure of only 200 pounds per square inch.

Instead of introducing the aluminum chloride as a slurry in a part of the charging stock We may introduce it as a slurry solution or fluent stream in previously formed catalyst complex. Thus the complex settled from warm settler H1 may vbe introduced through line |19 into slurry tank Si instead of supplying a slurrying liquid to this tank through line 9|. The aluminum 'chloride complex for dissolving or suspending the make-up aluminum chloride may be from any source whatsoever but it is preferably a complex prepared from a light paraiiinic hydrocarbon the light isoparafn complexes being particularly suitable. Complex may be withdrawn from reactor H91 `or reactor |46 to serve as a vehicle for introducing make-up aluminum chloride andJ in fact, any of the complex or sludge formed in the system may be used for this purpose if it is sufficiently nonviscous. The slurry or solution of aluminum chloride in complex may be made at a slightly higher temperature than the temperature prevailing in the reactor and suitable settling or agitating means may be employed in slurry tank s2 for eecting an intimate mixture of aluminum chloride with the complex. rI'his fluent stream of make-up aluminum chloride in complex may be injected into the reactor 'by means of suitable pumps or it may simply be introduced by fluid pressure. If gravity head is not sufficient, the pressure may be applied by huid means such for example as high pressure naphtha or hydrogen. .The make-up aluminum chloride may be supplied intermittently instead of continuously.

The plant hereinabove described is designed to convert a 152 F. end point parainic light naphtha into approrimately 98 volume percent yields of high octane number isomate of approximately the same boiling range. The end point of the total isomate may be somewhat higher than that of the charging stock but the 90 point of the product will be substantially the same as the 90% point of the charge. The lighter ends of the product may be slightly more volatile than the lighter ends of the charge. The Reid vapor pressure of the product may be slightly higher than that of the charge but the difference will not be substantially greater than about two pounds. With a charge of about 69 ASTM octane number the total isomate will have an octane number of about 80 or 8l and if the heavier components of the isomate are removed therefrom by fractionation the remainder will have an ASTM octane number within the approximate range of 85 to 90. l cc. of tetraethyl lead will give an increase of about l octane number to this product and 3 cc. of tetraethyl lead will give an increase of almost 20 in octane number.

While we have described in detail certain spe cic examples of our invention it should be understood that the invention is not limited to these specific examples nor to the specic apparatus or operating conditions described in connection therewith. Other types of apparatus, alternative methods of operation and equivalent operating conditions for accomplishing the objects of our invention will be apparent from the above description to those skilled in the art. In this specication and in the accompanying claims, the word barrel refers to a barrel of 42 U. S. gallon capacity. The cubic feet of hydrogen refers to the volume of hydrogen measured at 60 F. and 760 millimeters pressure. The Word continuously includes intermittent as well as uniform addition or withdrawal of materials. Aluminum chloride is intended to include the known aluminum halide equivalents of aluminum halide such as aluminum bromide, etc. Similarly, hydrogen chloride is intended to include equivalent hydrogen halides such as hydrogen bromide. Gas pressures are intended by simple reference to pressure and pounds per square inch. In other words, Applicants specification is addressed to those skilled in the petroleum refining art and is, therefore, couched in language familiar to those skilled in said art.

We claim:

1. In a process wherein paraffin hydrocarbons containing more than four but less than seven carbon atoms per molecule are isomerized by passage in admixture With hydrogen chloride under isomerization conditions through a reaction zone containing an aluminum chloride isomerization catalyst, and a fixed gas containing hydrogen and hydrogen chloride is separated from the reaction product, the combination of steps which comprises recycling at least a portion of said fixed gas and intimately contacting said recycled chloride absorption zone, introducing hydrogen chloride at a lower point of said absorption zone, employing a temperature in the absorption zone in the range of about 60 F. to about 100 F. and a pressure in the absorption Zone in the range of about v to about 200 pounds per square inch and such as to elect solution 'of a determinable amount of hydrogen chloride in the absorption zone, and introducing said hydrocarbons with dissolved hydrogen chloride to the reaction zone with a total amount of hydrogen chloride within the range of about 2 to 10% by weight based on total hydrocarbons charged thereto.

2. In a process wherein parafn hydrocarbons containing more than four but less than seven carbon atoms per molecule are isomerized by passage in admixture with hydrogen chloride under isomerization conditions through a reaction zone containing an aluminum chloride isomerization catalyst, and a ixed gas containing hydrogen and hydrogen chloride is separated from the reaction product, the combination of steps which comprises recycling at least a portion of said fixed gas and contacting said recycled fixed gas with said hydrocarbons prior to the introduction of said hydrocarbons into said reaction zone, initially introducing at least a part of said hydrocarbons at an upper part of a hydrogen chloride absorption zone, introducing hydrogen chloride at a lower part of said absorption zone, employing a pressure in the absorption Zone sucient to maintain hydrocarbons therein in liquid phase and sucient to eiect solution of a determinable amount of hydrogen chloride in the liquid hydrocarbon phase in the absorption zone, and introducing said hydrocarbons with dissolved hydrogen chloride to the reaction zone with a. total amount of hydrogen chloride within the range of about 2 to 10% by weight based on total hydrocarbons charged thereto.

3. The process of claim 2 wherein the isomerization conditions include a temperature in the range of about 200 F. to 350 F., a total operating pressure in the range of about 500 to 1500 pounds per square inch gauge and a space velocity of about 1.4 volumes of charging stock per hour per volume of catalyst material in the reactor.

BERNARD L. EVERING. EDMOND L. DOUV'ILLE.

REFERENCES CITED The following references are of record in the le of this patent:

UNITED STATES PATENTS Number Name Date 2,208,362 Engel July 16, 1940 2,220,091 Evering et al Nov. 5, 1940 2,250,118 Smithuysen July 22, 1942 2,266,012 dOuville et al Dec. 16, 1941 2,280,710 Lynch Apr. 21, 1942 2,300,249 Evering et al Oct. 27, 1942 2,314,297 Watson Mar. 16, 1943 FOREIGN PATENTS Number Country Date 24,044 India Aug. 23, 1937 

